Method and system embodiments for converting ethanol to para-xylene and ortho-xylene

ABSTRACT

Disclosed herein are embodiments of a method and system for converting ethanol to para-xylene. The method also provides a pathway to produce terephthalic acid from biomass-based feedstocks. In some embodiments, the disclosed method produces p-xylene with high selectivity over other aromatics typically produced in the conversion of ethanol to xylenes, such as m-xylene, ethyl benzene, benzene, toluene, and the like. And, in some embodiments, the method facilitates the ability to use ortho/para mixtures of methylbenzyaldehyde for preparing ortho/para xylene product mixtures that are amendable to fractionation to separate the para- and ortho-xylene products thereby providing a pure feedstock of para-xylene that can be used to form terephthalic anhydride and a pure feedstock of ortho-xylene that can be used for other purposes, such as phthalic anhydride.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a divisional of U.S. application Ser. No.17/387,725, filed Jul. 28, 2021; this prior application is incorporatedherein by reference in its entirety.

ACKNOWLEDGMENT OF GOVERNMENT SUPPORT

This invention was made with Government support under ContractDE-AC05-76RL01830 awarded by the U.S. Department of Energy. TheGovernment has certain rights in the invention.

FIELD

The present disclosure is directed to method embodiments for makingpara-xylene (or “p-xylene”) and ortho-xylene (or “o-xylene”) fromethanol, as well as system embodiments used for the method.

PARTIES TO JOINT RESEARCH AGREEMENT

The claimed invention arose under an agreement between Battelle MemorialInstitute and LanzaTech, Inc., which agreement was in effect on orbefore the effective filing date of the claimed invention.

BACKGROUND

Terephthalic acid is a high volume and high market commodity chemical.It is one of the compounds used to make polyethylene terephthalate,known as PET, which is used for making beverage bottles, packagingfilms, and fibers. Terephthalic acid is currently made from thepetroleum derived p-xylene produced from the naphtha reforming process.Potential unavailability of the petroleum-based p-xylene to meet themarket demand and the end user's interest in sustainable PET productshas created attention towards sustainable based feed source for theterephthalic acid production. Current ethanol-to-p-xylene processeseither require an excessive number of catalytic steps, or producep-xylene at low selectivity, thereby requiring capital-intensiveseparations. There exists a need in the art for a catalytic processusing renewably sourced ethanol to provide polymer-grade p-xylene thatcan serve as the basis for the economical and renewable terephthalicacid production.

SUMMARY

Disclosed herein are embodiments of a method, comprising: contacting afeed stream comprising ethanol with an oxidation catalyst underoxidation conditions to form an oxidation zone effluent streamcomprising acetaldehyde; passing the oxidation zone effluent stream to adimerization zone and contacting the oxidation zone effluent stream witha dimerization catalyst under dimerization conditions to produce adimerization zone effluent stream comprising 2-butenal; passing thedimerization zone effluent stream to a cyclization zone and contactingthe dimerization zone effluent stream with a cyclization catalyst undercyclization conditions to form a cyclization zone effluent streamcomprising o-methylbenzaldehyde and/or p-methylbenzaldehyde; and passingthe cyclization zone effluent stream to a hydrogenation zone andcontacting the cyclization zone effluent stream with a hydrogenationcatalyst comprising a first Group VIII metal deposited on a supportmaterial to produce a hydrogenation zone effluent comprising anon-equilibrium mixture of xylenes.

Also disclosed herein are embodiments of an apparatus comprising: a gasfermentation bioreactor in fluid communication with an oxidationreactor; the oxidation reactor in fluid communication with adimerization reactor; the dimerization reactor in fluid communicationwith a cyclization reactor; the cyclization reactor in fluidcommunication with a hydrogenation reactor; the hydrogenation reactor influid communication with a first fractionation zone; the firstfractionation zone in fluid communication with a second fractionationzone; and the second fractionation zone in fluid communication with afirst crystallizer.

The foregoing and other objects and features of the present disclosurewill become more apparent from the following detailed description, whichproceeds with reference to the accompanying figures.

BRIEF DESCRIPTION OF THE DRAWINGS

FIGS. 1A and 1B provide schematic diagrams outlining representativemethod embodiment steps and/or system components for converting ethanolto xylene compounds, such as o-xylene and p-xylene, wherein FIG. 1Asummarizes steps and system components used in certain disclosedembodiments and FIG. 1B summarizes steps used in certain disclosedembodiments.

FIG. 2 shows results obtained from evaluating carbon yield (%) for arepresentative method for making ortho-methylbenzaldehyde (or“o-methylbenzaldehyde”) and para-methylbenzaldehyde (or“p-methylbenzaldehyde”) products from acetaldehyde using (i) adimerization step in a dimerization zone to produce 2-butenal fromacetaldehyde and (ii) a cyclization step using a cyclization zone toproduce the ortho-methylbenzaldehyde and para-methylbenzaldehydeproducts from the 2-butenal made in the dimerization zone.

FIG. 3 shows results obtained from evaluating conversion and selectivity(%) for a cyclization step, wherein 2-butenal is reacted with a catalystat varying temperature and pressure; bar A summarizes reaction productdistribution for a cyclization step using a temperature of 260° C. and apressure of 101.35 kPa (14.7 psig); bar B summarizes reaction productdistribution for a cyclization step using a temperature of 150° C. and apressure of 689.5 kPa (100 psig); bar C summarizes reaction productdistribution for a cyclization step using a temperature of 175° C. and apressure of 1034 kPa (150 psig); and bar D summarizes reaction productdistribution for a cyclization step using a temperature of 200° C. and apressure of 1034 kPa (150 psig).

FIG. 4 shows conversion and total product yield results for condensingand cyclizing 2-butenal using fresh and regenerated TiO₂ catalyst.

FIG. 5 shows the product distribution obtained from condensing andcyclizing 2-butenal using fresh and regenerated TiO₂ catalyst.

FIGS. 6A and 6B show conversion and total product yield results forcondensing and cyclizing 2-butenal to methylbenzaldehyde (FIG. 6A), andthe corresponding product distribution obtained using differenthydrotalcite-based catalysts (FIG. 6B).

FIG. 7 shows conversion and total product yield results for condensingand cyclizing 2-butenal to methylbenzaldehyde using an Mg₄Al₁ catalystwith varying amounts of Na.

FIG. 8 shows conversion and total product yield results for condensingand cyclizing 2-butenal to methylbenzaldehyde using a Mg₄Al₁ catalystwith Na or K.

FIG. 9 shows conversion results obtained from evaluating the effect ofMg₄Al₁ catalyst regeneration on converting of 2-butenal tomethylbenzaldehyde.

FIG. 10 shows product selectivity results obtained from evaluating theeffect of Mg₄Al₁ catalyst regeneration on converting of 2-butenal tomethylbenzaldehyde.

FIG. 11 shows conversion and selectivity (%) results obtained for ahydrogenation step, wherein a mixture comprising ortho/paramethylbenzaldehyde was reacted with a Pd catalyst and a Re modifiercomponent at 180° C. and 6895 kPa (1000 psig) H₂ for six hours usingdifferent amounts of the Pd/Re catalyst system (i.e., 2 wt %, 2.9 wt %,5 wt %, and 10.6 wt %); “DMC” is dimethyl cyclohexane.

FIG. 12 shows conversion and selectivity (%) results obtained for ahydrogenation step using a batch combinatorial protocol, wherein amixture comprising ortho/para methylbenzaldehyde was reacted with a Pdcatalyst (with and without a Re modifier component) at different Pdloadings (i.e., 3 wt %, 0.75 wt %, and 0.25 wt %).

FIG. 13 shows conversion and selectivity (%) results obtained for ahydrogenation step using a batch combinatorial protocol, wherein amixture comprising ortho/para methylbenzaldehyde was reacted with a Pdcatalyst and a Re modifier component at different Pd:Re ratios wherein0.1 wt % Pd was used with varying amounts of Re.

FIG. 14 shows conversion and selectivity (%) results obtained for ahydrogenation step using a batch combinatorial protocol, wherein amixture comprising ortho/para methylbenzaldehyde was reacted with a Pdcatalyst and a Re modifier component at different Pd:Re ratios wherein0.25 wt % Pd was used with varying amounts of Re.

FIG. 15 shows conversion and selectivity (%) results obtained for ahydrogenation step using a flow reactor protocol, wherein a mixturecomprising ortho/para methylbenzaldehyde was reacted with a Pd catalystand a Re modifier component (3 wt % Pd and 6 wt % Re) on a carbonsupport.

FIG. 16 shows conversion and selectivity (%) results obtained for ahydrogenation step using a flow reactor protocol, wherein a mixturecomprising ortho/para methylbenzaldehyde was reacted with a Pd catalyst(0.25 wt %) without a Re modifier component on a carbon support.

FIG. 17 shows conversion and selectivity (%) results obtained for ahydrogenation step using a flow reactor protocol, wherein a mixturecomprising ortho/para methylbenzaldehyde was reacted with a Pd catalystand a Re modifier component (0.25 wt % Pd and 0.5 wt % Re) on a carbonsupport.

FIG. 18 shows conversion and selectivity (%) results obtained for ahydrogenation step using a flow reactor protocol, wherein a mixturecomprising ortho/para methylbenzaldehyde as reacted with a Pd catalystand a Re modifier component (0.1 wt % Pd and 0.2 wt % Re) on a carbonsupport.

FIG. 19 shows conversion and selectivity (%) results obtained for ahydrogenation step using a flow reactor protocol, wherein a mixturecomprising ortho/para methylbenzaldehyde was reacted with a Pd catalyst(0.25 wt %) without a Re modifier component on a carbon support for atleast 600 hours, time on stream.

FIG. 20 shows conversion and selectivity (%) results obtained for ahydrogenation step using a flow reactor protocol, wherein a mixturecomprising ortho/para methylbenzaldehyde was reacted with a Pd catalystand a Re modifier component (0.5 wt % Pd and 1 wt % Re) on a carbonsupport for at least 400 hours, time on stream.

DETAILED DESCRIPTION I. Overview of Terms

The following explanations of terms and abbreviations are provided tobetter describe the present disclosure and to guide those of ordinaryskill in the art in the practice of the present disclosure. As usedherein, “comprising” means “including” and the singular forms “a” or“an” or “the” include plural references unless the context clearlydictates otherwise. The term “or” refers to a single element of statedalternative elements or a combination of two or more elements, unlessthe context clearly indicates otherwise.

Unless explained otherwise, all technical and scientific terms usedherein have the same meaning as commonly understood to one of ordinaryskill in the art to which this disclosure belongs.

Although methods and materials similar or equivalent to those describedherein can be used in the practice or testing of the present disclosure,suitable methods and materials are described below. The materials,methods, and examples are illustrative only and not intended to belimiting. Other features of the disclosure are apparent from thefollowing detailed description and the claims.

Unless otherwise indicated, all numbers expressing quantities ofcomponents, molecular weights, molarities, voltages, capacities, and soforth, as used in the specification or claims are to be understood asbeing modified by the term “about.” Accordingly, unless otherwiseimplicitly or explicitly indicated, or unless the context is properlyunderstood by a person of ordinary skill in the art to have a moredefinitive construction, the numerical parameters set forth areapproximations that may depend on the desired properties sought and/orlimits of detection under standard test conditions/methods as known tothose of ordinary skill in the art. When directly and explicitlydistinguishing embodiments from discussed prior art, the embodimentnumbers are not approximates unless the word “about” is recited.

Although the operations of exemplary embodiments of the disclosed methodmay be described in a particular, sequential order for convenientpresentation, it should be understood that disclosed embodiments canencompass an order of operations other than the particular, sequentialorder disclosed, unless the context dictates otherwise. For example,operations described sequentially may in some cases be rearranged orperformed concurrently. Further, descriptions and disclosures providedin association with one particular embodiment are not limited to thatembodiment and may be applied to any disclosed embodiment.

Although there are alternatives for various components, parameters,operating conditions, etc. set forth herein, that does not mean thatthose alternatives are necessarily equivalent and/or perform equallywell. Nor does it mean that the alternatives are listed in a preferredorder unless stated otherwise.

In order to facilitate review of the various embodiments of thedisclosure, the following explanations of specific terms are provided:

Cyclization Catalyst: A catalyst that is capable of promoting 2-butenalcondensation and cyclization to o-methylbenzaldehyde and/orp-methylbenzaldehyde.

Cyclization Conditions: Reaction conditions, such as temperature,pressure, reaction time, weight hourly space velocity, and/orcyclization catalyst loading that can be controlled and/or modified tofacilitate 2-butenal condensation and cyclization too-methylbenzaldehyde and/or p-methylbenzaldehyde.

Cyclization Zone: A reaction zone comprising system componentsconfigured to contact a dimerization zone effluent stream comprising2-butenal with a cyclization catalyst to form a cyclization zoneeffluent stream comprising o-methylbenzaldehyde and/orp-methylbenzaldehyde

Dimerization Catalyst: A catalyst that is capable of promotingacetaldehyde dimerization to 2-butenal.

Dimerization Conditions: Reaction conditions, such as temperature,pressure, reaction time, weight hourly space velocity, and/ordimerization catalyst loading that can be controlled and/or modified tofacilitate acetaldehyde dimerization to 2-butenal.

Dimerization Zone: A reaction zone comprising system componentsconfigured to contact an oxidation zone effluent stream comprisingacetaldehyde with a dimerization catalyst to form a dimerization zoneeffluent stream comprising 2-butenal.

Feed Stream: A fluid stream that is passed to one or more zones. Anexemplary feed stream is a fluid stream comprising ethanol that can beintroduced into an oxidation zone.

Fractionization Zone: A zone comprising system components capable offractionating one fluid component from another (e.g., fractionatingp-xylene from o-xylene).

Hydrogenation Catalyst: A catalyst that is capable of promotingo-methylbenzaldehyde and/or p-methylbenzaldehyde hydrogenation to axylene product mixture, wherein the xylene product mixture compriseso-xylene and/or p-xylene. In particular embodiments, the xylene productmixture is a non-equilibrium mixture of xylenes.

Hydrogenation Conditions: Reaction conditions, such as temperature,pressure, reaction time, and/or hydrogenation catalyst loading that canbe controlled and/or modified to facilitate hydrogenation ofo-methylbenzaldehyde to o-xylene and/or p-methylbenzaldehyde top-xylene.

Hydrogenation Zone: A reaction zone comprising system componentsconfigured to contact a cyclization zone effluent stream comprisingo-methylbenzaldehyde and/or p-methylbenzaldehyde with a hydrogenationcatalyst to form an effluent stream comprising a xylene product mixture.In particular embodiments, the xylene product mixture is anon-equilibrium mixture of xylenes.

Non-Equilibrium Mixture of Xylenes: A mixture of xylene compounds,wherein the mixture comprises p-xylene, o-xylene, and meta-xylene (or“m-xylene”, wherein the concentration of any m-xylene in the mixture isless than 50 wt % of a m-xylene equilibrium concentration. The mixtureof xylenes is exclusive of other compounds containing eight (8) carbonatoms.

Oxidation Catalyst: A catalyst that is used to promote convertingethanol to acetaldehyde.

Oxidation Conditions: Reaction conditions, such as temperature,pressure, reaction time, and/or oxidation catalyst loading that can becontrolled and/or modified to facilitate converting ethanol toacetaldehyde.

Oxidation Zone: A reaction zone comprising system components configuredto contact a feed stream comprising ethanol with an oxidation catalystto form an oxidation zone effluent stream comprising acetaldehyde.

II. Introduction

Polymer-grade p-xylene is a valuable product in various industries suchas the production of terephthalic acid, which in turn is used to producevarious polymers. Polymer-grade p-xylene used in these industries needsto have a purity of at least 99.95 mass % p-xylene, or at least 99.97(or greater) mass % p-xylene. Current processes which yield high purityp-xylene with sufficient commercial yield require substantial investmentin purification and isomerization operations including vessels, recycleof effluent streams, and utilities, all of which have high capital andoperating expenditures. Further, methods that exist in the art toproduce terephthalic acid from ethanol rely on oxidizing apara-methylbenzaldehyde product formed during the process to theterephthalic acid. Such methods typically require focusing on theability to increase the amount of para-methylbenzaldehyde produced inthe process in order to arrive at a sufficient amount of the material tobe oxidized to the terephthalic acid. This can typically require usingexpensive catalysts and/or processing parameters that do not lend toindustrial usage.

The disclosure herein provides embodiments of a method to provide highpurity p-xylene, such as polymer-grade p-xylene, without the need forisomerizing large amounts of undesired xylene products, such asm-xylene, and subsequent recycling. Additionally, the disclosuredescribes method embodiments that use separation techniques, such ascrystallization, which is less costly as compared to adsorptiveseparation. Furthermore, the disclosed method embodiments are compatiblewith unconventional feedstocks, such as ethanol, which may be derivedfrom a sustainable source. In some embodiments, the sustainable sourceof ethanol may be industrial waste gases, such as steel mill gas, orsyngas from various sources such as gasification of biomass ormunicipal/industrial waste. In some embodiments the sustainable sourceof ethanol may be a gas comprising CO₂.

In particular embodiments, a method for making p-xylene from ethanol isdisclosed, which provides a novel method for producing p-xylene,including polymer-grade p-xylene, and, in some embodiments, terephthalicacid from sustainable-based feedstocks. In some embodiments, thedisclosed method produces p-xylene with high selectivity over otheraromatics, such as m-xylene, ethyl benzene, benzene, toluene, and thelike. The disclosed method is more efficient and more economical thanconventional methods. Also, the method can produce mixtures of o-xyleneand p-xylene from methylbenzyaldehyde that can be fractionated toseparate the p- and o-xylene products, thereby providing a stream highlyenriched in p-xylene. This stream can be introduced to an additionalpurification process to economically produce polymer-grade p-xylene andan enriched stream of o-xylene that can be used for other purposes, suchas phthalic anhydride production. Parameters of the novel method (e.g.,reagents and/or reaction conditions) can be controlled to provide p/omethylbenzaldehyde mixtures and p/o-xylene product mixtures that includelittle to no undesired products, such as undesired aromatics (e.g.,m-xylene, toluene, or benzene), and/or saturated cyclic products (e.g.,dimethylcyclohexane). These are just a few of the improvements that canbe achieved using the method embodiments disclosed herein.

III. Method Embodiments

The present disclosure describes embodiments of a method for producingp-xylene from ethanol. Ethanol used in the disclosed method can beobtained from petroleum-derived ethanol from ethylene, or ethanolderived from a sustainable source, such as industrial waste or offgases, such as steel mill gas, syngas from various sources (e.g.,gasification of biomass or municipal/industrial waste), or gascomprising CO₂. In some embodiments, p-xylene produced using thedisclosed method can be used to produce terephthalic acid. In yet someadditional embodiments, o-xylene produced by the method can be furtherconverted to phthalic anhydride.

In particular embodiments of the method, a feed stream comprisingethanol and, optionally, an oxygen stream or a stream containing asource of oxygen, is contacted with an oxidation catalyst underoxidation conditions in an oxidation zone to form an effluent streamcomprising acetaldehyde. The oxidation zone can comprise a reactor orvessel that contains the oxidation catalyst. Alternatively, theoxidation zone can be one section of a reactor or vessel. If theoxidation zone is one section of a reactor, the reactor may comprise oneor more reaction zones as described hereinafter. The oxidation catalystcan be present in the reactor as a stationary bed through which the feedstream is flowed through or it can be present as a moving bed or asparticulates that are fluidized and flowed co-currently or countercurrently with the feed stream. The vessel or reactor can comprise oneor more inlets for introducing the reactants and one or more outlets forremoving the products and unreacted reactants. In some embodiments, theethanol is mixed with air and contacted with the catalyst at atemperature ranging from 500° C. to 650° C. One oxidation process knownas Verb Chemie Process involves mixing ethanol with air prior tointroducing the mixture into the oxidation vessel or zone or the ethanoland air can be introduced as separate streams into the oxidationzone/vessel. Such a process can be used in the current method. In someembodiments, the acetaldehyde can be formed from ethanol bydehydrogenation processes. Such processes will utilize catalystscomprising a metal oxide or carbon support material or zeolite materialand a metal or metals selected from Cu, Au, Ni, Zn, Mn, Co, V, Ag, Fe,Ce, or Cr. In particular embodiments, the catalyst is selected from acatalyst comprising copper and chromium and either a mesoporous carbonsupport material or a Al₂O₃ support material.

The feed stream for the oxidation zone can comprise ethanol that isderived from C1 gas fermentation of a source material. The fermentationcan use the source material directly such as in the fermentation ofcellulosic material, or indirectly such as through the gasification ofbiomass to produce syngas. Examples of source material includecellulosic material, sugars, industrial process waste gas or non-wastegas, combustion engine exhaust fumes, such as automobile exhaust fumes,biogas, landfill gas, direct air capture, from electrolysis orcombinations thereof. The substrate and/or C1-carbon source of the gasfermentation to generate the ethanol feed stream may be syngas generatedby pyrolysis, torrefaction, or gasification. In other words, carbon inwaste material may be recycled by pyrolysis, torrefaction, orgasification to generate syngas which is used as the substrate and/orC1-carbon source in the gas fermentation that generates the ethanol feedstream. The substrate and/or C1-carbon source in the gas fermentationmay be a gas stream comprising methane.

In particular embodiments, the feed stream can comprise ethanol derivedfrom waste gas produced by an industrial process selected from ferrousmetal products manufacturing, steel mill manufacturing, non-ferrousproducts manufacturing, petroleum refining, electric power production,carbon black production, paper and pulp production, ammonia production,methanol production, coke manufacturing, petrochemical production,carbohydrate fermentation, cellulosic fermentation, cement making,aerobic digestion, anerobic digestion, catalytic processes, natural gasextraction, oil extraction, geological reservoirs, gas from fossilresources such as natural gas coal and oil, or any combination thereof.Examples of specific processing steps within an industrial processinclude catalyst regeneration, fluid catalyst cracking, and catalystregeneration. Air separation and direct air capture are other suitableindustrial processes. In these embodiments, the substrate and/orC1-carbon source for the gas fermentation to generate the ethanol in thefeed stream may be captured from the industrial process before it isemitted into the atmosphere, using any known method.

In yet some additional embodiments, the feed stream can comprise ethanolderived from synthesis gas, known as syngas, which may be obtained frompyrolysis, torrefaction, reforming, partial oxidation, or gasificationprocesses. Examples of gasification processes include gasification ofcoal, gasification of refinery residues, gasification of petroleum coke,gasification of biomass, gasification of lignocellulosic material,gasification of waste wood, gasification of black liquor, gasificationof municipal solid waste, gasification of municipal liquid waste,gasification of industrial solid waste, gasification of industrialliquid waste, gasification of refuse derived fuel, gasification ofsewerage, gasification of sewerage sludge, gasification of sludge fromwastewater treatment, gasification of biogas. Examples of reformingprocesses include, steam methane reforming, steam naphtha reforming,reforming of natural gas, reforming of biogas, reforming of landfillgas, naphtha reforming, and dry methane reforming. Examples of partialoxidation processes include thermal and catalytic partial oxidationprocesses, catalytic partial oxidation of natural gas, partial oxidationof hydrocarbons. Examples of municipal solid waste include tires,plastics, fibers, such as in shoes, apparel, and textiles. Municipalsolid waste may be simply landfill-type waste. The municipal solid wastemay be sorted or unsorted. Examples of biomass may includelignocellulosic material and may also include microbial biomass.Lignocellulosic material may include agriculture waste and forest waste.

The substrate and/or C1-carbon source may be a gas stream comprisingmethane. Such a methane containing gas may be obtained from fossilmethane emission such as during fracking, wastewater treatment,livestock, agriculture, and municipal solid waste landfills. It is alsoenvisioned that the methane may be burned to produce electricity orheat, and the C1 byproducts may be used as the substrate or carbonsource.

The fermentation of gaseous stream comprising C1 compounds (e.g., CO,CO₂, CH₄, CH₃OH, etc.) to produce products such as ethanol and aceticacid are well known in the art. The fermentation process comprisescontacting a gaseous C1-containing stream with a at least one C1-fixingbacteria in a liquid medium in a bioreactor. In one embodiment, thebacteria can be selected from the genus Clostridia. Exemplaryfermentation processes are described in U.S. Pat. Nos. 8,507,228;8,263,372; 8,809,015; and 8,663,949, the relevant portions of which areincorporated herein by reference.

As discussed above, in some embodiments, the feed stream may compriseethanol that is derived from liquid fermentation of sugars and orcellulosic material. In yet additional embodiments, the feed stream maycomprise ethanol from hydration of ethylene. The ethanol may also beproduced from traditional ethanol manufacturing processes, or in otherwords, ethanol from a source other than cellulosic material, sugar,industrial process waste gas, automobile exhaust fumes, or syngas fromgasification operations. In particular embodiments, the oxidation zonecan convert at least 20 wt % of the ethanol of the feed stream toacetaldehyde, such as 20 wt % to 95 wt % of the ethanol, or 50 wt % to90 wt % of the ethanol, or 70 wt % to 90 wt % of the ethanol. Theoxidation zone operates at temperatures from 200° C. to 500° C. Inparticular embodiments, the oxidation zone operates 250° C. to 400° C.In some embodiments, the oxidation zone can be operated under suitablepressures, which are recognized by those of ordinary skill in the artwith the benefit of the present disclosure.

The effluent stream from the oxidation zone comprising acetaldehyde isnext contacted with a dimerization catalyst to produce an effluentstream comprising 2-butenal. This further involves passing the effluentstream comprising acetaldehyde from the oxidation zone to thedimerization zone. The dimerization zone can comprise a separate reactoror vessel, or a separate section of the reactor or vessel, that housesthe oxidation zone. The dimerization catalyst can be present in thereactor as a stationary bed through which the acetaldehyde stream isflowed through or can be present as particulates that are fluidized andflowed co-currently or counter currently with the acetaldehyde stream.The vessel or reactor can comprise one or more inlets for introducingthe reactants and one or more outlets for removing the products andunreacted reactants. If the dimerization zone is housed in the samereactor or vessel as the oxidation zone, then the effluent from theoxidation zone can be passed directly to the dimerization zone. Thedimerization zone is operated under dimerization conditions.Dimerization conditions can comprise a reaction temperature ranging from150° C. to 310° C., such as 160° C. to 300° C., or 180° C. to 300° C.and can be conducted at pressures ranging from 689.5 kPa (100 psig) to1034 kPa (150 psig), such as 689.5 kPa (100 psig), 758 kPa (110 psig),or 1034 kPa (150 psig). In some embodiments, the dimerization catalystis a catalyst that comprises an oxide material, such as MnO, MgO, ZnO,ZrO₂, TiO₂, or any combination thereof. In some embodiments, thecatalyst can further comprise a support material, such as an aluminasupport (e.g., Al₂O₃). In particular embodiments, the catalyst isselected from MnO—ZnO—ZrO₂, MgO—Al₂O₃, ZnO—ZrO₂ (10:1), ZnO—ZrO₂ (2:1),or TiO₂. In particular embodiments, the dimerization zone can convert atleast 15 wt % of the acetaldehyde of the effluent produced by theoxidation zone to a product mixture comprising 2-butenal, such as 15 wt% to 65 wt % or higher of the acetaldehyde, or 20 wt % to 65 wt % of theacetaldehyde, or 32 wt % to 65 wt %, or 59 wt % to 65 wt % or higher ofthe acetaldehyde. In some embodiments, 58 wt % or more of the productmixture can be 2-butenal, such as 68 wt % to 91 wt %, or 77 wt % to 91wt %, or 81 wt % to 91 wt %, or 84 wt % to 91 wt %, or 88 wt % to 91 wt%.

In some embodiments, exposing the effluent stream comprising 2-butenalto a cyclization catalyst to produce an effluent stream comprising amixture of o-methylbenzaldehyde and/or p-methylbenzaldehyde can comprisecontacting the effluent stream comprising the 2-butenal with thecyclization catalyst in a cyclization zone. This further involvespassing the effluent stream comprising 2-butenal from the dimerizationzone to a cyclization zone. The cyclization zone can comprise a separatereactor or vessel, or a separate section of the reactor or vessel, thathouses the dimerization zone or the oxidation zone plus the dimerizationzone. The cyclization catalyst can be present in the reactor as astationary bed through which the 2-butenal stream is flowed through, orcan be present as particulates that are fluidized and flowedco-currently or counter currently with the acetaldehyde stream. Thevessel or reactor can comprise one or more inlets for introducing thereactants and one or more outlets for removing the products andunreacted reactants. If the cyclization zone is housed in the samereactor or vessel as the dimerization zone, then the effluent from thedimerization zone can be passed directly to the cyclization zone. Thecyclization zone can be operated under cyclization conditions.Cyclization conditions can comprise using a reaction temperature rangingfrom 250° C. or higher, such as 250° C. to 350° C., or 250° C. to 325°C., or 275° C. to 300° C. In particular embodiments, the cyclizationstep is performed at pressures ranging from atmospheric pressure 101.35kPa (14.7 psig) to 1034 kPa (150 psig), such as atmospheric pressure to689.5 kPa (100 psig). In some embodiments, a weight hourly spacevelocity (or “WHSV”) ranging from 0.2 h⁻¹ to 0.25 h⁻¹, such as 0.2 h⁻¹to 0.23 h⁻¹, or 0.2 h⁻¹ to 0.22 h⁻¹ is used. In particular embodiments,the cyclization zone can be operated at 150° C. and 689.5 kPa (100psig), or at 260° C. and 101.35 kPa (14.7 psig), or at 175° C. and 1034kPa (150 psig), or at 200° C. and 1034 kPa (150 psig), or at 300° C. and101.35 kPa (14.7 psig).

In some embodiments, the cyclization catalyst is a metal oxide catalyst,such as a metal oxide catalyst comprising a Group IV metal (also knownas Group 4 under the new IUPAC classification), such as Ti or Zr; or aGroup II metal (also known as Group 2 under the new IUPACclassification), such as Mg.

In some independent embodiments, the cyclization catalyst can furthercomprise a solid support, such as an alumina support (e.g., Al₂O₃), SrO,CaO, MgO, La₂O₃, SiO₂, SiO₂—Al₂O₃, TiO₂—Al₂O₃, or a zeolite support suchas H-Mordenite, or Faujasite, onto which the desired metal oxide isdispersed or deposited. In some embodiments, the cyclization catalyst isselected from TiO₂ or a hydrotalcite catalyst comprising MgO and Al₂O₃.In particular embodiments, the hydrotalcite catalyst can have a formulaof Mg_(x)Al_(y), wherein “Mg” represents MgO, “Al” represents Al₂O₃, xranges from 1 to 4, and y typically is 1. In some such embodiments, thehydrotalcite catalyst can be Mg₂Al₁, Mg₃Al₁, or Mg₄Al₁. In someembodiments, the hydrotalcite catalyst can be modified to comprise analkali metal, such as Na or K. The amount of alkali metal included inthe catalyst can range from greater than 0 wt % to 20 wt % or higher,such as greater than 0 wt % to 20 wt %, or greater than 0 wt % to 10 wt%, or greater than 0 wt % to 5 wt %. In particular embodiments, thecatalyst used in the cyclization zone can be selected from TiO₂; Mg₄Al₁;Mg₃Al₁; Mg₂Al₁; Mg₄Al₁ comprising 5 wt % or 10 wt % or 20 wt % Na; orMg₄Al₁ comprising 5 wt % K. In particular embodiments, the cyclizationzone can convert at least 50 wt % of the 2-butenal of the effluentproduced by the dimerization zone to a resulting product mixturecomprising o-methylbenzaldehyde and p-methylbenzaldehyde, such as 50 wt% to 95 wt % or higher of the 2-butenal, or 75 wt % to 95 wt % of the2-butenal, or 80 wt % to 95 wt % of the 2-butenal. In particularembodiments, the dimerization step provides a product mixture comprisingo-methylbenzaldehyde and p-methylbenzaldehyde but nometa-methylbenzaldehyde (or “m-methylbenzaldehyde”). In some suchembodiments, the reaction mixture may comprise trace amounts (e.g., lessthan 20 wt % total) of other compounds, such as higher aldehydes (e.g.,2,4,6-octatrienal), benzaldehyde, or hydrogenated products. In someembodiments, the reaction mixture may only comprise other compounds(e.g., higher aldehydes, benzaldehyde, and/or hydrogenated products) inan amount ranging from greater than 0% to 20%, greater than 0% to 15%,or greater than 0% to 10% of the reaction mixture.

In some embodiments, the effluent stream comprising o-methylbenzaldehydeand/or p-methylbenzaldehyde is contacted with a hydrogenation catalystin a hydrogenation zone to produce an effluent comprising a xyleneproduct mixture. This involves passing the effluent stream comprisingthe o-methylbenzaldehyde and/or p-methylbenzaldehyde mixture from thecyclization zone to the hydrogenation zone. The hydrogenation zonecomprises a reactor or vessel in which the hydrogenation catalyst iscontained. In particular embodiments, the hydrogenation zone comprises areactor vessel suitable for flow-based processing) comprising one ormore inlets and outlets and that is configured to contain the catalyst.The hydrogenation zone can comprise a separate reactor or vessel, or aseparate section of the reactor or vessel, that houses the hydrogenationzone or any or all of the previous zones plus the hydrogenation zone.The hydrogenation catalyst can be present in the reactor as a stationarybed through which the cyclization zone effluent stream is flowed throughor can be present as particulates that are fluidized and flowedco-currently or counter currently with the cyclization zone effluentstream. The vessel or reactor can comprise one or more inlets forintroducing the reactants and one or more outlets for removing theproducts and unreacted reactants. If the hydrogenation zone is housed inthe same reactor or vessel as the cyclization zone, then the effluentfrom the cyclization zone can be passed directly to the hydrogenationzone. The hydrogenation zone can be operated under hydrogenationconditions. Hydrogenation conditions can comprise operating thehydrogenation zone at temperatures ranging from greater than 100° C. toless than 200° C., such as 110° C. to 190° C., or 120° C. to 180° C., or125° C. to 175° C. In particular embodiments, the temperature is 125°C., 130° C., 140° C., 150° C., 160° C., 170° C., or 180° C., or anytemperatures between 125° C. and 180° C. In some embodiments,hydrogenation conditions can comprise operating the hydrogenation zoneat pressures ranging from greater than 344.7 kPa (50 psig) to less than1378.9 kPa (2000 psig), such as 689.5 kPa (100 psig) to 10342 kPa (1500psig), or 1378.9 kPa (200 psig) to 6894.7 kPa (1000 psig), or 3447.4 kPa(500 psig) to 5515.8 kPa (800 psig). In particular embodiments, thepressure is 689.5 kPa (100 psig), 1034 kPa (150 psig), 1378.9 kPa (200psig), 1723.7 kPa (250 psig), 2068.4 kPa (300 psig), 2413.2 kPa (350psig), 2757.9 kPa (400 psig), 3102.6 kPa (450 psig), 3447.4 kPa (500psig), 3792 kPa (550 psig), 4136.8 kPa (600 psig), 4481.6 kPa (650psig), 4826.3 kPa (700 psig), 5171.1 kPa (750 psig), 5515.8 kPa (800psig), 5860.5 kPa (850 psig), 6205.3 kPa (900 psig), 6550 kPa (950psig), or 6894.7 kPa (1000 psig). In particular embodiments, thehydrogenation zone is operated at 125° C. and 689.5 kPa (100 psig), 150°C. and 3447.4 kPa (500 psig), at 150° C. and 6894.7 kPa (1000 psig), or180° C. and 6894.7 kPa (1000 psig). In particular embodiments, thehydrogenation zone is operated at temperatures ranging from 125° C. to175° C. or lower and at pressures ranging from 689.5 kPa (100 psig) to5515.8 kPa (800 psig) or less.

In some embodiments, hydrogenation conditions can comprise operating thehydrogenation zone at any of the above temperatures and/or pressures fora time period sufficient to convert all, or substantially all, of theo-methylbenzaldehyde and/or p-methylbenzaldehyde to the correspondingxylene product mixture. In such embodiments, “substantially all” meansat least 90 wt %, such as at least 93 wt %, or at least 94 wt %, or atleast 95 wt %, or at least 96 wt %, or at least 97 wt %, or at least 98wt %, or at least 99 wt % of the effluent comprising theo-methylbenzaldehyde and/or p-methylbenzaldehyde. In some embodiments,the time period ranges from minutes to an amount of time whereby theprocess is stopped or the catalyst efficiency decreases by 50%. In someembodiments, the time period can range from 30 minutes to 600 hours ormore, such as 1 hour to 600 hours (or more), or 3 hours to 600 hours (ormore), or 6 hours to 600 hours (or more). In particular embodiments, thetime period can be 30 minutes, 1 hour, 6 hours, 400 hours, and 600hours. In particular embodiments, the hydrogenation zone can convert atleast 85 wt % of the mixture of the o-methylbenzaldehyde andp-methylbenzaldehyde of the effluent produced by the cyclization zone tothe xylene product mixture, such as 85 wt % to 100 wt % of the mixtureof the o-methylbenzaldehyde and p-methylbenzaldehyde, or 90 wt % to 100wt % of the mixture of the o-methylbenzaldehyde andp-methylbenzaldehyde, or 95 wt % to 100 wt % of the mixture of theo-methylbenzaldehyde and p-methylbenzaldehyde.

In some embodiments, the hydrogenation catalyst comprises a Group VIII(also known as Group 8, 9, or 10 under the new IUPAC classification)metal and a support material. In particular embodiments, the Group VIIImetal of the hydrogenation catalyst is selected from iron, ruthenium,cobalt, rhodium, iridium, nickel, palladium, platinum, or combinationsthereof. In particular embodiments, the Group VIII metal is palladium,platinum, or ruthenium. In representative embodiments, the Group VIIImetal is palladium. The support material of the hydrogenation catalystcan be selected from carbon, silicas, aluminas, silica-aluminas,titania, zirconia, zeolites, zinc oxides, or combinations thereof. Inparticular embodiments, the support material is a carbon material, asilica, an alumina (e.g., Al₂O₃), a titania (e.g., TiO₂), a zirconia(e.g., ZrO₂), a niobium oxide (e.g., Nb₂O₅), a low acidity zeolite(e.g., ZSM-5), or a carbon support (a carbon support selected from acarbon support sold under the tradenames NucharO sold by Ingevity, aHyperion 07C or Hyperion 02C support sold by Hyperion CatalysisInternational, ROX HF or DARCO-LS supports sold by Cabot Norit, a carbonsupport sold by Jacobi, CECA, or PICA). The Group VIII metal typicallyis deposited on the support material by means known to those in the artand with the benefit of the present disclosure and exists in metallicform on the support under reaction conditions disclosed herein. Inparticular embodiments, the support material is carbon, Al₂O₃, TiO₂, orZrO₂. In some embodiments, the support material can be substantiallynon-acidic. Acidity of the support material can be measured bydetermining the surface concentration of acid sites of the supportmaterial using Fourier transform infra-red (FTIR) analysis of adsorbedpyridine. The amount of the Group VIII metal used in the catalyst,expressed as the amount of metal per total weight of the catalyst, canrange from 0.1 wt % to 3 wt %, such as 0.5 wt % to 2.5 wt %, or 0.75 wt% to 2 wt %, or 1 wt % to 2 wt %. In some embodiments, the amount of theGroup VIII metal used can range from 0.75 wt % to 3 wt %, such as 1.5 wt% to 3 wt % per total weight of the catalyst. In some other embodiments,the amount of the Group VIII metal used can range from greater than 0 wt% to less than 0.75 wt %, such as greater than 0 wt % to 0.5 wt %, orgreater than 0 wt % to 0.25 wt %, or greater than 0 wt % to 0.2 wt %, orgreater than 0 wt % to 0.15 wt %, or greater than 0 wt % to 0.1 wt % pertotal weight of the catalyst. In particular embodiments, the amount ofthe Group VIII metal can be used in an amount selected from 0.1 wt %,0.25 wt %, 0.5 wt %, 0.75 wt %, 1.5 wt %, or 3 wt % per total weight ofthe catalyst. In an independent embodiment, the hydrogenation catalystcomprises Rainey nickel, without a support material.

In some embodiments, the hydrogenation catalyst can further comprise amodifier component. The modifier component can be an element, typicallya metal, that is used in combination with a metal catalyst to modifyproperties of the catalyst. In some embodiments, the modifier componentcan facilitate stabilizing the metal catalyst and/or reducing acidicsites on a support material. In particular embodiments, the modifiercomponent is a metal that is deposited on the support of thehydrogenation catalyst. In some embodiments, the modifier component cancomprise a Group VII metal (also known as Group 7 under the new IUPACclassification), or a Group IV metal (also known as Group 14 under thenew IUPAC classification), a Group I metal (also known as Group 1 underthe new IUPAC classification and/or as an alkali metal), a Group IImetal (also known as Group 2 under the new IUPAC classification and/oras an alkaline earth metal), or a combination thereof. In particularembodiments, the modifier component can comprise rhenium, tin, orcombinations thereof. The amount of the modifier component used in thecatalyst, expressed as the amount of metal per total weight of catalyst,can range from 0 wt % to 6 wt %, such as 0.1 wt % to 6 wt %, or 0.2 wt %to 4 wt %, or 0.5 wt % to 3 wt %, or 1.5 wt % to 2 wt %. In someembodiments, the amount of the modifier component used in the catalystcan range from 3 wt % to 6 wt %, such as 5 wt % to 6 wt %, expressed asthe amount of modifier component per total catalyst. In some otherembodiments, the amount of the modifier component used in the catalystcan range from 0.1 wt % to 1 wt % such as 0.1 wt % to 0.5 wt %, or 0.1wt % to 0.2 wt % expressed as the amount of modifier component per totalcatalyst. In particular embodiments, the amount of the modifiercomponent present on the catalyst is selected from 0.1 wt %, 0.2 wt %,0.5 wt %, 1 wt %, 1.5 wt %, 5 wt %, or 6 wt %, expressed as the amountof modifier component per total catalyst. In particular embodiments, thehydrogenation zone can further comprise a solvent. In some embodiments,the solvent can be hydrocarbon solvent, such as dodecane, decane, anddioxane.

In representative embodiments, the hydrogenation catalyst comprises theGroup VIII metal, the support, and the modifier component. In yet otherembodiments, the hydrogenation catalyst comprises the Group VIII metaland the support. In particular embodiments, the hydrogenation catalystconsists of the Group VIII metal and the support. In yet additionalparticular embodiments, the hydrogenation catalyst consists of the GroupVIII metal, the support, and the modifier component. In an independentembodiment, the hydrogenation catalyst consists of Rainey nickel. Thecomponents of the hydrogenation catalyst may be added sequentially inany order, or in any combination, including all together at the sametime. In representative embodiments, the hydrogenation catalystcomprises palladium, a carbon material, and rhenium; or, palladium,Al₂O₃, and rhenium; or palladium, TiO₂, and rhenium; or palladium, ZrO₂,and rhenium; platinum, a carbon material, and rhenium; or, platinum,Al₂O₃, and rhenium; or platinum, TiO₂, and rhenium; or platinum, ZrO₂,and rhenium; or ruthenium, a carbon material, and rhenium; or,ruthenium, Al₂O₃, and rhenium; or ruthenium, TiO₂, and rhenium; orruthenium, ZrO₂, and rhenium. In yet other representative embodiments,the hydrogenation catalyst comprises palladium and Al₂O₃; or palladiumand ZrO₂; or palladium and TiO₂, or palladium and carbon; or platinumand Al₂O₃; or platinum and ZrO₂; or platinum and TiO₂, or platinum andcarbon; or ruthenium and Al₂O₃; or ruthenium and ZrO₂; or ruthenium andTiO₂, or ruthenium and carbon. In some embodiments, the method cancomprise using 0.1 wt % to 3 w % of the catalyst. In particularembodiments, such amounts correspond to (i) the total amount of theGroup VIII metal, the support material, and any modifier component, ifpresent, and is relative to the feed; or (ii) the total amount of themetal relative to the feed, such as in embodiments using Rainey nickelwithout a support material. In particular embodiments, 2 wt % of thecatalyst is used. In particular representative embodiments, thehydrogenation catalyst comprises a mixture of 3 wt % Pd and 6 wt % Re ona carbon support material. In yet other embodiments, the hydrogenationcatalyst comprises 1.5 wt % Pd and 3 wt % Re on a carbon supportmaterial; or 0.5 wt % Pd and 1 wt % Re on a carbon support material; or0.1 wt % Pd and 2 wt % Re on a carbon support material; or 0.25 wt % Pdand 0.5 wt % Re on a carbon support material; or 0.25 wt % Pd and 0.5 wt% Re on a carbon support material; or 0.25 wt % Pd and 1.5 wt % Re on acarbon support material; or 0.25 wt % Pd and 6 wt % Re on a carbonsupport material; or 0.1 wt % Pd and 0.5 wt % Re on a carbon supportmaterial; or 0.1 wt % Pd and 1.5 wt % Re on a carbon support material;or 0.1 wt % Pd and 6 wt % Re on a carbon support material; or 0.75 wt %Pd and 5 wt % Re on a carbon support material; or 3 wt % Pd and 6 wt %Re on Al₂O₃; or 1.5 wt % Pd and 3 wt % Re on Al₂O₃; or 0.75 wt % Pd and5 wt % Re on Al₂O₃; or 1.5 wt % Pd on Al₂O₃; or 1.5 wt % Pd on ZrO₂; or0.25 wt % Pd on a carbon material.

In some particular embodiments, the xylene product mixture in thehydrogenation zone effluent comprises at least 85 wt % of a mixturecomprising o-xylene and/or p-xylene. The balance of the xylene productmixture in the hydrogenation zone effluent may include toluene, benzene,m-xylene, ethylbenzene, and/or saturated aromatic compounds, such asdimethylcyclohexane. The hydrogenation zone effluent typically does notcomprise an equilibrium amount of ethylbenzene. In some particularembodiments, the xylene product mixture consists essentially of o-xyleneand p-xylene, wherein consisting essentially of means that the xyleneproduct mixture is substantially free of an isomerized version ofo-xylene and/or p-xylene (e.g., m-xylene), a saturated aromatic compound(e.g., dimethylcyclohexane), and/or cracked aromatic compounds (e.g.,toluene or benzene) such that the amount of any such products,individually, is less than 15 wt %; or such that the amount of any suchproducts, in total, is less than 30 wt %. Amounts of components presentin the hydrogenation zone effluent can be determined using standardizedtechniques and methods recognizable by those of skill in the art withthe benefit of the present disclosure. One exemplary method is gaschromatography coupled with flame ionization detection calibrated withexternal standards, with compounds being identified by massspectroscopy. In some embodiments, the xylene product mixture cancomprise o-xylene and/or p-xylene, and substantially no m-xylene. In anindependent embodiment, the xylene product mixture consists of o-xyleneand p-xylene.

In some embodiments, the xylene product mixture comprises p-xylene at aconcentration ranging from 65 wt % to 100 wt %, such as at least 65 wt%, or at least 75 wt %, or at least 85 wt %. In such embodiments, anyremaining weight balance can be o-xylenes or a mixture of o-xylenes andtrace amounts (e.g., less than 5 wt % total) of other aromatics (e.g.,m-xylene, benzene, or toluene) or a saturated aromatic (e.g.,dimethylcyclohexane). In some embodiments, the p-xylene concentrationranges from 65 wt % to 99 wt %, or 70 wt % to 99 wt %, or 75 wt % to 99wt %, or 80 wt % to 99 wt %, or 85 wt % to 99 wt %. In some otherembodiments, the xylene product mixture comprises o-xylene, p-xylene,and m-xylene. In such embodiments, the xylene product mixture comprisesa non-equilibrium mixture of the three different xylene products. Insuch non-equilibrium mixtures, m-xylenes is present at a concentrationlower than 50 wt % of a m-xylene equilibrium concentration. In someembodiments, if any isomerization to m-methylbenzaldehyde takes placeand any such m-methylbenzaldehyde is converted to m-xylenes, theresulting amount of m-xylenes is less than 50% of a m-xylene equilibriumconcentration. In some embodiments, if m-xylene is present, it ispresent at a concentration ranging from greater than 0 wt % up to 49 wt% of a m-xylene equilibrium concentration, such as greater than 0 wt %up to 45 wt %, or greater than 0 wt % to 40 wt %, or greater than 0 wt %to 35 wt %, greater than 0 wt % to 30 wt %, greater than 0 wt % to 20 wt%, greater than 0 wt % to 10 wt %, greater than 0 wt % to 5 wt %, orgreater than 0 wt % to 1 wt % of a m-xylene equilibrium concentration.

In some embodiments, the method can further comprise converting thepurified p-xylene to terephthalic acid. The purified p-xylene can beconverted to terephthalic acid under oxidative conditions that would berecognized by those of skill in the art with the benefit of the presentdisclosure. In some embodiments, this conversion can be conducted in airusing acetic acid with a manganese or cobalt acetate catalyst. In yetadditional embodiments, the method can further comprise converting theterephthalic acid to Polyethylene Terephthalate (PET). In suchembodiments, the method can comprise combining the terephthalic acidwith ethylene glycol in an esterification reactor or vessel, which canbe operated under conditions known to those of skill in the art with thebenefit of the present disclosure, such as at pressures of 206.8 kPa (30psig) to 344.7 kPa (50 psig) and temperatures ranging from 230° C. to260° C. Vapors produced during the method (e.g., water/steam and glycol)can be vented to a reflux column or distillation column and recoveredand returned to the esterification vessel/reactor or zone (in the caseof the glycol by-product) or discharged to waste (in the case ofcondensed water produced from steam condensation). A monomer is producedfrom this step, namely bis-(2-hydroxyethyl)-terephthalate (or “BHET”),which can be delivered to a second esterification and/or apolymerization reactor or zone wherein the BHET is polymerized to PET.

In some embodiments, the method can further comprise converting anyo-xylenes of the xylene product mixture to phthalic anhydride. In someembodiments, converting the o-xylene to phthalic anhydride can compriseprocessing the o-xylene at oxidation conditions known to those of skillin the art with the benefit of the present disclosure. The o-xylene canbe processed at oxidation conditions when present in an effluentcomprising the xylene product mixture, or a stream comprising o-xylenecan first be separated from the effluent comprising the xylene productmixture. In embodiments where a stream of the o-xylene is separated fromthe hydrogenation zone effluent comprising the xylene product mixture,the hydrogenation zone effluent is passed to a fractionation zonewherein the stream of o-xylenes is isolated from the hydrogenation zoneeffluent. The fractionation zone can comprise a separation column orother component suitable for fractionation through which thehydrogenation zone effluent comprising the xylene product mixture ispassed. Fractionation techniques suitable for this step and/or otherfractionation steps contemplated herein are described in more detailbelow. In some embodiments, the method can further comprise drying theeffluent comprising the xylene product mixture. In some suchembodiments, the hydrogenation zone effluent can be dried prior topassing it to the fractionation zone. In some additional embodiments,the method can further comprise drying the stream of o-xylene. In someembodiments, the method can further comprise drying both thehydrogenation zone effluent comprising the xylene product mixture andthe stream of o-xylene.

Fractionation is a commonly used method for many processes in manyindustrial plants to separate chemicals. In the present disclosure, afirst fractionation zone may be used to separate saturated cycliccompounds, such as dimethylcyclohexane, and some lighter compounds froma C8 aromatics stream. Unlike other technologies, the disclosed methodembodiments do not produce high volumes of C9 aromatics, so afractionation column that is typically known as a “xylene column,” whichis a large and costly fractionation column to separate C9 aromatics fromC8 aromatics, is not necessary. The fractionation column of the firstfractionation zone is far smaller with fewer theoretical stages since C9aromatics are not present in appreciable amounts and do not need to beseparated from the C8 aromatics. The overall cost, including both thecapital cost and the operating cost, of the first fractionation zone isdramatically reduced compared to systems where C9 aromatics need to beseparated from C8 aromatics.

The C8 aromatics from the first fractionation zone are then passed to asecond fractionation zone. It is often difficult to use conventionalfractional distillation technology to separate different xylene isomersand ethylbenzene, if present, efficiently and economically because theboiling points of such C8 aromatics fall within a very narrow 8° C.range, from about 136° C. to about 144° C. (see Table 1). The boilingpoints of p-xylene and ethylbenzene are about 2° C. apart, and theboiling points of p-xylene and m-xylene are only about 1° C. apart. As aresult, large equipment, significant energy consumption, and substantialrecycles are typically required to provide effective and satisfactoryxylene separations.

TABLE 1 C8 Compound Boiling Point (° C.) Freezing Point (° C.)Ethylbenzene 136 −95 p-xylene 138 +13 m-xylene 139 −48 o-xylene 144 −25

However, due to the unique hydrogenation zone effluent having anon-equilibrium mixture of xylenes, and specifically, having lower thanequilibrium amounts of m-xylene, the fractionation in the secondfractionation zone may be successfully accomplished by less costlyequipment. It is not necessary for the equipment to separate thep-xylene and m-xylene because the amount of m-xylene in thehydrogenation zone effluent is not an equilibrium amount of m-xylene,and in many embodiments will be substantially less than an equilibriumamount of m-xylene. For example, it is accepted that an equilibriumreaction for the conversion of toluene to xylenes and benzene productsnormally provides m-xylene in an amount from 64 mol % at −23.2° C. to 51mol % at 276.9° C., such as 62 mol % at −23.2° C. to 53 mol % at 276.9°C., or 60 mol % at −23.2° C. to 55 mol % at 276.9° C., or 58 mol % at−23.2° C. to 56 mol % at 276.9° C. In contrast, the hydrogenation zoneeffluent obtained with the presently disclosed method may comprise lessthan 50 wt % of the m-xylene equilibrium concentration, such as fromgreater than 0 wt % up to 49 wt % or greater than 0 wt % up to 45 wt %,or greater than 0 wt % to 40 wt %, or greater than 0 wt % to 35 wt %,greater than 0 wt % to 30 wt %, greater than 0 wt % to 20 wt %, greaterthan 0 wt % to 10 wt %, greater than 0 wt % to 5 wt %, or greater than 0wt % to 1 wt % of the m-xylene equilibrium concentration.

Furthermore, it is not necessary to achieve a pure p-xylene stream fromthe second fractionation zone and one effluent of the secondfractionation zone may contain a mixed xylene stream enriched inp-xylene, o-xylene, and trace m-xylene. With partial separation from thesecond fractionation zone being acceptable for the success of thedisclosure, further cost savings are achieved. In one embodiment, thesecond fractionation zone is designed to achieve chemical grade o-xylenein the overhead, and once that is achieved, the remainder of allcomponents may be removed in the bottoms. Thus, in some embodiments, asecond effluent from the second fractionation zone can contain a streamenriched to 99 wt % o-xylene, which may be collected and passed forother uses as discussed herein.

In some embodiments, the mixed xylene steam from the secondfractionation zone is passed to a crystallizer. Crystallizers can beused to purify the mixed xylene stream to polymer grade p-xylene. Insome embodiments, the mixed xylene stream feed to the crystallizer maycomprise less than 70 wt % p-xylene, which may benefit from one or morestages of crystallizers to produce polymer grade p-xylene at acceptablerecoveries. In one embodiment, the mixed xylene stream resulting fromfractionation will be 70 wt % or greater p-xylene, which can provideacceptable recoveries of polymer grade p-xylene using a single stagecrystallizer. In some embodiments, 99.5 wt % or 99.8 wt % p-xylene canbe obtained. In particular embodiments, at least 99.5 wt % or at least99.8 wt % p-xylene is obtained.

Crystallizers take advantage of the differences between the freezingpoints and solubilities of the C8 aromatic components at differenttemperatures. With its higher freezing point, p-xylene is usuallyseparated as a solid, while the other components are recovered in ap-xylene depleted filtrate. Crystallization results in polymer-gradepurity p-xylene, which typically is needed for commercial conversion ofp-xylene to terephthalic acid. Suitable crystallization processes aredescribed in U.S. Pat. Nos. 4,120,911 and 3,662,013, the relevantportions of which are incorporated herein by reference, and componentsused in such methods are commercially available.

In some embodiments, a first and second fractionation zone that isreduced in size and utility cost, combined with one or morecrystallizers, is a cost effective approach that can be used with thedisclosed method to achieve polymer grade p-xylene that capitalizes onthe unique advantages afforded by the composition of the hydrogenationzone effluent of the present disclosure.

In any or all of the described embodiments, the method can furthercomprise separating an unreacted reactant from an effluent producedduring the method and recycling the unreacted reactant to the zone fromwhich it was obtained and/or an upstream zone (wherein “upstream” isintended to indicate one or more previous zones relative to the zonefrom which the unreacted reactant is obtained). Solely by way ofexample, unreacted ethanol can be recycled from the oxidation zone bypassing any unreacted ethanol back into the oxidation zone, either viaan independent inlet of the reactor or container of the oxidation zone,or by recombining the unreacted ethanol with the feed stream and addingthe mixture into a feed stream inlet of the reactor or container. Suchrecycling can be used to increase the yield of acetaldehyde produced bythe oxidation zone. In yet other embodiments, any unreacted acetaldehydefrom the dimerization zone can be recycled back to the dimerization zoneso as to increase the amount of 2-butenal in the effluent produced fromthe dimerization zone. In such embodiments, the unreacted acetaldehydecan be recycled back into the dimerization zone either via anindependent inlet of the reactor or container of that zone, or byrecombining it with the effluent comprising acetaldehyde produced by theoxidation zone. In yet some additional embodiments, any unreacted2-butenal from the cyclization zone can be recycled back to thecyclization zone to increase the amount of o-methylbenzaldehyde andp-methylbenzaldehyde produced by the cyclization zone. In suchembodiments, the unreacted 2-butenal can be recycled back into thecyclization zone either via an independent inlet of the reactor orcontainer of that zone, or by recombining it with the effluentcomprising 2-butenal produced by the dimerization zone. In yet someadditional embodiments, any unreacted o-methylbenzaldehyde and/orp-methylbenzaldehyde from the hydrogenation zone can be recycled back tothe hydrogenation zone so as to increase the amount of the p-xyleneand/or o-xylene in the xylene product mixture produced by thehydrogenation zone. In such embodiments, the unreactedo-methylbenzaldehyde and/or p-methylbenzaldehyde can be recycled backinto the hydrogenation zone either via an independent inlet of thereactor or container of that zone, or by recombining it with theeffluent comprising o-methylbenzaldehyde and p-methylbenzaldehydeproduced by the cyclization zone. Any combination of these recyclingembodiments can be used.

Steps and components of a representative method and system embodimentare summarized schematically in FIG. 1A. Representative method steps offurther embodiments are summarized schematically in FIG. 1B. By way ofexample, FIG. 1A shows a flow scheme of one embodiment of the presentdisclosure. Syngas or industrial gas in line 104 and, optionally,hydrogen in line 102, are passed to gas fermentation zone 100 having atleast one gas fermentation bioreactor comprising at least one C1-fixingbacteria in a liquid nutrient medium to generate gas fermentation zoneeffluent 106 comprising ethanol. Gas fermentation zone effluent 106comprising ethanol is passed to oxidation zone 110 where it is contactedwith an oxidation catalyst under oxidation conditions and produceoxidation zone effluent 108 comprising acetaldehyde, which in turn ispassed to dimerization zone 120 where it is contacted with adimerization catalyst under dimerization conditions and produce adimerization zone effluent 112 comprising 2-butenal. Dimerization zoneeffluent 112 is passed to cyclization zone 130 where it is contactedwith a cyclization catalyst under cyclization conditions and producecyclization zone effluent 114 comprising o-methylbenzaldehyde and/orp-methylbenzaldehyde, which in turn is passed to hydrogenation zone 140where it is contacted with a hydrogenation catalyst comprising a firstgroup VIII metal (IUPAC 8, 9, and 10) optionally deposited on a supportmaterial under hydrogenation conditions to produce hydrogenation zoneeffluent 116 comprising a xylene product mixture, which comprises anon-equilibrium mixture of xylenes.

Hydrogenation zone effluent 116 is passed to first fractionation zone150 where benzene-enriched stream is removed in overhead stream 118 anda dimethylcycohexane-enriched stream is removed in a bottoms stream 122.The remainder of hydrogenation zone effluent 116 is removed in stream124 and passed to second fractionation zone 160. The amount of benzenemay influence the point at which stream 124 is removed from the firstfractionation zone, thus FIG. 1A shows a generic location, and does notindicate a sidecut per se. Chemical grade purity o-xylene stream 126 isremoved from second fractionation zone as a bottoms stream, and xyleneproduct stream 128 comprising p-xylene, o-xylene, and trace m-xylene isremoved from second fractionation zone as an overhead stream.

Xylene product stream 128 is passed to crystallizer 170. Depending uponthe composition of xylene product stream 128, crystallizer 170 may becontrolled such that p-xylene product stream 134 has sufficiently highpurity to meet polymer grade purity standards. The o-xylene filtratestream 132 has a low enough content of p-xylene such that when o-xylenefiltrate stream 132 is combined with stream 126 from the secondfractionation zone, the combined stream still meets chemical gradepurity levels for o-xylene. p-Xylene product stream 134 which is astream of polymer grade purity p-xylene may be the final desired productstream of the process. Furthermore, as shown, p-xylene product streammay be derived from a source of recycled carbon as shown in FIG. 1A.Optionally, p-xylene product stream 134, or p-xylene product stream 138discussed below, may be passed to a catalytic liquid phase oxidationreactor 190 for the conversation of the p-xylene to terephthalic acid,which is then removed in stream 144.

Shown in FIG. 1A is an optional second crystallizer 180. When xyleneproduct stream 128 contains less than 70 wt % p-xylene, a secondcrystallizer 180 may be used. In this embodiment, the p-xylene productstream 134 from crystallizer 170 is passed to second crystallizer 180 togenerate p-xylene product stream 138 having polymer grade p-xylene ando-xylene stream 136 having a sufficiently low amount of p-xylene so thatafter combining with stream 132 to form combined stream 137, and furthercombining with stream 126, the resulting combined stream 142 still meetschemical grade purity levels of o-xylene.

Another benefit of the flow scheme illustrated by FIG. 1A is that thecrystallizers which are used are much smaller than those used inconventional production of polymer grade p-xylene. Small scalecrystallizers may facilitate the ability to verify the purified polymergrade p-xylene obtained from recycled carbon or from a sustainablesource. Some recycled carbon or sustainable sources of C1 substrates forgas fermentation to produce ethanol provide C1 substrates on a smallscale and small scale crystallizers provide the ability to carry-out theprocess from the generation of ethanol by gas fermentation throughpurification of p-xylene on a scale commensurate with the C1 substratesupply for the gas fermentation which is useful to verify and certifythe p-xylene is sustainable or derived from recycled carbon.

Products obtained from method embodiments disclosed herein can be usedin various applications and techniques to make additional products suchas articles of manufacture. In some embodiments, PET made according to amethod embodiment of the present disclosure can be converted intovarious PET-based products or articles, such as containers (bottles,jars, cans, coolers, etc.), packaging materials (food containers,storage containers, etc.), fibers (e.g., threads and yarns for use infabrics and textiles), and films (wrapping materials, liners, foodwraps, etc.). In some embodiments, the PET material disclosed herein canbe converted to such products using molding techniques suitable for PETprocessing. In some embodiments, the PET can be blow molded into aproduct using an extrusion or injection blow molding process. Inextrusion blow molding, a parison of the PET material is placed in amold and hot air is blown into the parison to inflate it into the formof the mold. The object is cooled, the mold opened, and the objectejected. In injection blow molding, the PET material is injection moldedinto a heated cavity, onto a core pin. The cavity mold forms the outershape of the part and is based off a core rod which shapes the inside ofa preform. The preform mold is opened and compressed air is injectedinto the preform and the object is blown, cooled, and then ejected. Insome other embodiments, an object can be made from the PET materialusing a thermoforming technique, wherein a sheet of the PET material isheated to a temperature below its melting point to achieve a glassy orsoft state and it is then stretched to contours of a mold. The materialis then cut with a die to provide the desired formed object. In yetadditional embodiments, melt spinning techniques can be used to make PETfibers, wherein the PET material (in the form of chips, granules, or thelike) is melted to form a solution and then forced through holes of aspinneret, after which fibers of the material are drawn (stretched) toprovide a fiber of a desired diameter.

IV. Overview of Several Embodiments

Disclosed herein are embodiments of a method, comprising: contacting afeed stream comprising ethanol with an oxidation catalyst underoxidation conditions to form an oxidation zone effluent streamcomprising acetaldehyde; passing the oxidation zone effluent stream to adimerization zone and contacting the oxidation zone effluent stream witha dimerization catalyst under dimerization conditions to produce adimerization zone effluent stream comprising 2-butenal; passing thedimerization zone effluent stream to a cyclization zone and contactingthe dimerization zone effluent stream with a cyclization catalyst undercyclization conditions to form a cyclization zone effluent streamcomprising o-methylbenzaldehyde and/or p-methylbenzaldehyde; and passingthe cyclization zone effluent stream to a hydrogenation zone andcontacting the cyclization zone effluent stream with a hydrogenationcatalyst comprising a first Group VIII metal deposited on a supportmaterial to produce a hydrogenation zone effluent comprising anon-equilibrium mixture of xylenes.

In any or all embodiments of the method, the hydrogenation catalystfurther comprises a second Group VIII metal, a modifier component, or acombination thereof, all deposited on the support material wherein thesecond Group VIII metal is not the same as the first Group VIII metal.

In any or all of the above embodiments, the modifier component isselected from rhenium, tin, an alkali metal, an alkali earth metal, orany combination thereof.

In any or all of the above embodiments, the hydrogenation catalystcomprises the modifier component and wherein the support material iscarbon, the first Group VIII metal is palladium, and the modifiercomponent is rhenium.

In any or all of the above embodiments, the support material is selectedfrom carbon material, a silica, an alumina, a silica-alumina, a titania,a zirconia, a zeolite, a zinc oxide, or any combination thereof.

In any or all of the above embodiments, the non-equilibrium mixture ofxylenes comprises m-xylene in an amount ranging from 0 wt % to less than40 wt % of a m-xylene equilibrium concentration.

In any or all of the above embodiments, the non-equilibrium mixture ofxylenes comprises m-xylene in an amount ranging from 0 wt % to 20 wt %of a m-xylene equilibrium concentration.

In any or all of the above embodiments, the non-equilibrium mixture ofxylenes comprises m-xylene in an amount ranging from 0 wt % to 5 wt % ofa m-xylene equilibrium concentration.

In any or all of the above embodiments, the non-equilibrium mixture ofxylenes comprises m-xylene in an amount ranging from 0 wt % to 1 wt % ofa m-xylene equilibrium concentration.

In any or all of the above embodiments, the ethanol is (i) ethanol fromliquid phase fermentation of cellulosic material and or sugar; (ii)ethanol from gas phase fermentation of industrial process waste ornon-waste gas, internal combustion engine exhaust fumes, syngas, directair capture, electrolysis, CO₂-containing gas or any combinationthereof; (iii) ethanol from a source other than cellulosic material,sugar, industrial process waste or non-waste gas, internal combustionengine exhaust fumes, gasification processes, syngas, direct aircapture, electrolysis, or CO₂-containing gas; or (iv) ethanol fromhydration of ethylene; or any combination of (i), (ii), (iii), and/or(iv).

In any or all of the above embodiments, the industrial process isselected from ferrous metal products manufacturing, steel millmanufacturing, non-ferrous products manufacturing, petroleum refining,electric power production, carbon black production, paper and pulpproduction, ammonia production, methanol production, coke manufacturing,petrochemical production, carbohydrate fermentation, cellulosicfermentation, cement making, aerobic digestion, anerobic digestion,catalytic processes, natural gas extraction, oil extraction or anycombination thereof; and/or wherein the syngas is from coalgasification, refinery residues gasification, petroleum cokegasification, biomass gasification, lignocellulosic materialgasification, waste wood gasification, black liquor gasification,natural gas reforming, municipal solid or liquid waste gasification,refuse derived fuel gasification, sewerage or sewerage sludgegasification, sludge from waste water treatment gasification and/orindustrial solid waste gasification or any combination thereof.

In any or all of the above embodiments, the conversion of acetaldehydein the dimerization zone provides 15 wt % to 65 wt % of a productreaction mixture comprising 2-butenal; the selectivity of acetaldehydeto 2-butenal in the dimerization zone ranges from 57 wt % to 91 wt %;the conversion of 2-butenal in the cyclization zone provides 70 wt % to95 wt % of a product reaction mixture comprising o-methylbenzaldehydeand p-methylbenzaldehyde; the selectivity of 2-butenal too-methylbenzaldehyde and p-methylbenzaldehyde in the cyclization zoneranges from 50 wt % to 95 wt %; or any combination of any of theaforementioned.

In any or all of the above embodiments, the method further comprisespassing the hydrogenation zone effluent to a fractionation zone andseparating a stream comprising o-xylene from (i) a stream comprisingp-xylene or (ii) a stream comprising p-xylene and m-xylene.

In any or all of the above embodiments, (i) the stream comprisingp-xylene or (ii) the stream comprising p-xylene and m-xylene comprises aminimum amount of p-xylene, wherein the minimum amount of p-xyleneranges from a minimum of at least 65 wt % to a minimum of at least 85 wt%.

In any or all of the above embodiments, the method further comprises (i)drying the stream comprising the o-xylene; (ii) reacting the o-xylene inthe stream comprising o-xylene under reaction conditions to formphthalic anhydride; or both (i) and (ii).

In any or all of the above embodiments, the method further comprisesdrying the hydrogenation zone effluent prior to passing it to thefractionation zone, and/or drying the stream comprising the o-xylene.

In any or all of the above embodiments, the method further comprisespassing (i) the stream comprising p-xylene or (ii) the stream comprisingp-xylene and m-xylene to a crystallizer and recovering a purifiedp-xylene stream comprising at least 99.5 wt % p-xylene.

In any or all of the above embodiments, the purified p-xylene streamcomprises at least 99.8 wt % p-xylene.

In any or all of the above embodiments, the method further comprisesreacting at least a portion of the p-xylene from the purified p-xylenestream under reaction conditions to form terephthalic acid.

In any or all of the above embodiments, the method further comprisesreacting at least a portion of the terephthalic acid with ethyleneglycol under reaction conditions to form polyethylene terephthalate.

In any or all of the above embodiments, the method further comprisesforming the polyethylene terephthalate into one or more products.

In any or all of the above embodiments, the method further comprises oneor more separation and/or recycling steps, wherein the recycling stepsare selected from (i) recycling at least a portion of the oxidation zoneeffluent stream to the oxidation zone until a predetermined targetconcentration of acetaldehyde in the oxidation zone effluent stream isachieved; (ii) recycling at least a portion of the dimerization zoneeffluent stream to the dimerization zone until a predetermined targetconcentration of 2-butenal in the dimerization zone effluent stream isachieved; (iii) recycling at least a portion of the cyclization zoneeffluent stream to the cyclization zone until a predetermined targetconcentration of o-methylbenzaldehyde and/or p-methylbenzaldehyde in thecyclization zone effluent stream is achieved; (iv) recycling at least aportion of the hydrogenation zone effluent stream to the hydrogenationzone until a predetermined target concentration of xylenes in thehydrogenation zone effluent stream is achieved; and/or (v) anycombination of steps (i), (ii), (iii), and/or (iv).

In any or all of the above embodiments, the method further comprisesregenerating the cyclization catalyst by heating the cyclizationcatalyst under air.

Also disclosed herein are embodiments of an apparatus, comprising: a gasfermentation bioreactor in fluid communication with an oxidationreactor; the oxidation reactor in fluid communication with adimerization reactor; the dimerization reactor in fluid communicationwith a cyclization reactor; the cyclization reactor in fluidcommunication with a hydrogenation reactor; the hydrogenation reactor influid communication with a first fractionation zone; the firstfractionation zone in fluid communication with a second fractionationzone; and the second fractionation zone in fluid communication with afirst crystallizer.

In any or all of the above embodiments, the apparatus further comprisesa second crystallizer in fluid communication with the firstcrystallizer.

In any or all of the above embodiments, the apparatus further comprisesa catalytic liquid phase oxidation reactor in fluid communication withthe first crystallizer.

In any or all of the above embodiments, the apparatus further comprisesa catalytic liquid phase oxidation reactor in fluid communication withthe second crystallizer.

V. Examples

General Procedure for Batch High Throughput Method used in Examples9-15: Powdered catalysts were weighed out into 2 mL glass reactionvials, performed in triplicate. 48 vials were assembled onto one highthroughput plate, which was then sealed and reduced under 5% H₂/N2reduction gas, ramping at 2° C./minute to 300° C. and holding for 4hours. The sealed reactor was transferred to a flow-through N2 purgebox, where it was unsealed and each vial was filled with 1.75 mL ofreactant solution with a composition of 12.5 wt % o-methylbenzaldehyde,12.5 wt % p-methylbenzaldehyde and remainder dodecane solvent. Thereactor was resealed and transferred out of the purge box and connectedto an automated batch reactor setup. Run operation began with 3 cyclesof reactor pressurization to 100 psi H₂ to purge out air in the lines,before flowing pure H₂ until the desired pressure is reached. All lineswere sealed so that the reactor was isolated, with 48 individual vialssharing the headspace. The reactor was heated at 4° C./minute whilebeing shaken in a circular motion at 600 RPM to facilitate masstransfer. After reaching target temperature and then holding for thelength of the experiment, the reactor was cooled down to roomtemperature before the pressure was released. GC-FID analysis was usedto quantitate the products.

Example 1

In this example, a representative dimerization reaction that takes placein a dimerization zone as described herein was evaluated. The feedcomposition used in this particular example was ethanol-derivedacetaldehyde. Results of particular examples are shown in FIG. 2 . And,data from particular examples are provided in Table 2.

TABLE 2 Catalyst Conditions Conversion Aldehyde Selectivity (%)Composition T (° C.) P (psig) P (kPa) (%) 2-butenal C4+ Total*MnO—ZnO—ZrO₂ 180 150 1034 64.8% 90.8% 96.9% 300 110 758 58.9% 83.9%95.4% MgO—Al₂O₃ 180 100 689.5 20.0% 81.1% 96.8% 300 110 758 60.8% 57.9%82.9% ZnO—ZrO₂ (10:1) 180 110 758 24.1% 87.5% 92.0% ZnO—ZrO₂ (2:1) 180110 758 32.0% 90.0% 95.9% TiO₂ 180 110 758 15.1% 77.4% 94.8% *C4+ Total:Sum of 2-butenal (C4), hexadienal (C6), octatrienal (C8)

Example 2

In this example, a representative cyclization reaction that takes placein a cyclization zone as described herein was evaluated. The feedcomposition used in this particular example was 2-butenal. Results froma particular example are shown in FIG. 3 . FIG. 4 also shows resultsfrom an example wherein both the 2-butenal conversion and thecorresponding total product yield obtained during condensation reactionat 300° C. using a TiO₂ catalyst were obtained. In this example, theTiO₂ catalyst surface area was 60 m²/g, the reaction temperature was300° C. (conducted at atmospheric pressure), and the WHSV was 0.2 h⁻¹.Very high 2-butenal conversion (˜95%) was achieved with a fresh TiO₂catalyst. In some examples, decreased conversion, along with the time onstream, suggested catalyst deactivation. While the conversion wasdecreased from ˜95% to ˜70% for some examples, the yield of the totalproduct remained constant (˜25%). The products that were detected byGC-MS in some examples included 2,4,6 octatrienal, o-methylbenzaldehydeand p-methylbenzaldehyde, o-xylene and benzaldehyde. Without beinglimited to a single theory, it currently is believed that the low carbonbalance obtained in some examples is attributed to the formation of longchain oligomeric products that are not detectable in the GC-MS.Effective regeneration of the TiO₂ catalyst was also demonstrated—seeFIG. 4 . The 2-butenal conversion was very similar as was obtained inthe fresh catalyst. However, the yield of the products was ˜15%, whichcurrently is believed to suggest that although a regenerated catalystprovides similar 2-butenal conversion, it also favors the formation oflong chain oligomeric products. In some examples, the TiO₂ catalystobtained after regeneration was further deactivated after time onstream, however, the total yield of the products did not changesignificantly.

Product distribution yield results obtained during an example of the2-butenal condensation/cyclization step is shown in FIG. 5 .Condensation and cyclization of C4, 2-butenal, yielded cyclic C8products, such as p-methylbenzaldehyde and o-methylbenzaldehyde, and anacyclic C8 product, such as 2,4,6 octatrienal. Among these products,o-methylbenzaldehyde was obtained as the major product. A very similartrend in product distribution was obtained with both fresh andregenerated catalysts, although the overall yield was higher in the caseof fresh catalyst. Without being limited to a single theory, itcurrently is believed that the formation of o-xylene might be attributedto o-methylbenzaldehyde reduction to the corresponding benzyl alcohol,followed by hydrodeoxygenation. The Bronsted acidity present in the TiO₂catalyst may catalyze the hydrodeoxygenation reaction at highertemperature.

Example 3

In this example, different hydrotalcite-based catalysts were evaluatedfor the condensation of 2-butenal. The results obtained from thesescatalysts at different process condition are shown in FIGS. 6A and 6B.In these examples, the reaction conditions were a temperature of 300°C., atmospheric pressure, and a WHSV of 0.23 h⁻¹. All the hydrotalcitecatalysts (MgO/Al₂O₃) showed higher activity compared to the MgOcatalyst as evidenced by the higher conversion of 2-butenal as shown inFIG. 6A. Only ˜10% conversion was achieved with the MgO catalyst;however, hydrotalcite-based catalysts showed ˜50% conversion. The totalproduct yield obtained in theses catalyst was ˜35%, further suggestinggood carbon balance using these catalysts. Thus, although hydrotalcitecatalysts provide lower conversion compared to TiO₂ catalysts, it isbelieved that they can prevent formation of unwanted long chainoligomeric products. FIG. 6A also shows the correlation between theeffect of Al content in the hydrotalcite catalyst and the correspondingactivity for 2-butenal condensation. Although the Mg₄Al₁ catalyst showeda 3-fold increase in conversion compared to the MgO catalyst, furtherincrease in Al content (i.e., Mg₄Al₁→Mg₃Al₁→Mg₂Al₁) did not show muchimpact as both conversion and product yield were similar in the exampleswith higher Al content. FIG. 6B shows the corresponding productdistribution obtained with different catalysts. The products in FIG. 6Bincluded 2,4,6 octatrienal, 2-methylbenzaldehyde (oro-methylbenzaldehyde), 2-methyl benzyl alcohol, 4-methyl benzaldehyde(or p-methylbenzaldehyde), 4-methyl benzyl alcohol, and benzaldehyde.

Example 4

In this example, hydrotalcite-based catalysts comprising differentamounts of impregnated Na were evaluated for the condensation of2-butenal. Mg₄Al₁+x % Na (x=0-20%) was prepared by impregnating Mg₄Al₁catalyst with different amount of Na. FIG. 7 shows the conversion andyield of the products obtained with Mg₄Al₁ catalyst containing varyingNa amounts. In this example, the reaction conditions were a temperatureof 300° C. at atmospheric pressure and a WHSV of 0.23 h⁻¹. While theincorporation of different amounts of Na did not have significant impacton conversion of 2-butenal, the overall product yield was increased withincreasing Na content, up to 10 wt %.

Example 5

In this example, condensation of 2-butenal was investigated using aMg₄Al₁ catalyst containing different alkali metals at a 5 wt. % loading.FIG. 8 shows the conversion and yield of the products obtained with theMg₄Al₁+5 wt. % M (where M is Na or K) catalysts versus a catalystcontaining only Mg₄Al₁. The reaction conditions were a temperature of300° C. at atmospheric pressure and a WHSV of 0.22 h⁻¹. The catalystscontaining an alkali metal showed both increased conversion and productyield compared to pure Mg₄Al₁ catalyst with K being more active than Na.

Example 6

In this example, the effects of Mg₄Al₁ catalyst regeneration onperformance was evaluated. FIG. 9 shows regeneration of Mg₄Al₁ catalystand consequent effect on catalytic performance. The reaction conditionswere a temperature of 300° C. at atmospheric pressure and a WHSV of 0.22h⁻¹. After the catalyst activity had decreased to 80% conversion of2-butenal, it was regenerated at 550° C. for 2 hours under air. As shownin FIG. 9 , the catalytic activity was regained after a firstregeneration, as both 2-butenal conversion product yield were verysimilar to those of the fresh catalyst. The first regenerated catalystwas again tested for condensation of 2-butenal until conversion haddropped to about 70%. The catalyst was again regenerated under the aboveconditions and its conversion and products yield were again restored toabout the same performance as the fresh catalyst. The second regeneratedcatalyst was again tested for condensation of 2-butenal until conversionhad dropped to about 80%. The catalyst was regenerated a third timeunder the above conditions and its conversion and product yield wasagain restored to about the same performance as the fresh catalyst.Using three regenerations, the catalyst life was extended to over 400hours. FIG. 10 shows the product selectivity was improved slightly withthe regenerated catalyst.

Example 7

In this example, a cyclization reaction that takes place in ahydrogenation zone as described herein was evaluated. A batch reactionwas conducted using reaction conditions of: 180° C., 6894.7 kPa (1000psig) H₂, and a reaction time of 6 hours. Different catalyst loadings ona carbon support material were used, namely 2 wt %, 2.9 wt %, 5 wt %,and 10.6 wt % (expressed as the amount of metal per total weight ofcatalyst) of a 3 wt % Pd/6 wt % Re catalyst mixture. Results are shownin FIG. 11 .

Example 8

In this example, different catalysts were evaluated for use in ahydrogenation reaction that takes place in a hydrogenation zone asdescribed herein. The feed composition used in this particular examplewas a mixture of 50% p-methylbenzaldehyde and 50% o-methylbenzaldehydein dodecane. The feed to solvent ratio was 1 to 4. The reactiontemperature was 150° C. and the reaction pressure was 3447.4 kPa (500psig). The reaction was run for a time period of 30 minutes. Conversionrates and carbon yield obtained using eight different hydrogenationcatalyst compositions are presented in Table 3.

TABLE 3 Carbon Yield, % Conversion, % Methylbenzyl SI CatalystMethylbenzaldehyde Alcohol Xylene # Composition Para Ortho Para OrthoPara Ortho 1 3 wt % Pd/6 100.0 100.0 0.0 0.0 50.0 50.0 wt % Re on Carbon2 1.5 wt % Pd/3 100.0 100.0 0.0 0.0 50.0 50.0 wt % Re on Carbon 3 0.75wt % Pd/5 100.0 100.0 0.0 0.0 49.5 50.0 wt % Re on Carbon 4 3 wt % Pd/6100.0 100.0 0.0 0.0 48.3 50.0 wt % Re on Al₂O₃ 5 1.5 wt % Pd/3 100.0100.0 0.0 0.0 38.8 42.4 wt % Re on Al₂O₃ 6 0.75 wt % Pd/5 100.0 100.032.1 28.0 19.2 22.4 wt % Re on Al₂O₃ 7 1.5 wt % Pd 100.0 100.0 24.1 20.824.0 26.2 on Al₂O₃ 8 1.5 wt % Pd 94.9 95.5 25.3 20.1 18.9 21.8 on ZrO₂

Example 9

In this example, combinatorial batch processes were evaluated using 3 wt% Pd 6 wt % Re/Hyperion C with the catalyst present at 2 wt % of thefeed. Temperature, pressure, and reaction time were evaluated, withresults presented in Table 4.

TABLE 4 Reaction Xylene Dimethyl Cyclohexane Methyl Benzyl PressurePressure Time Conversion Selectivity Selectivity (%) Alcohol SelectivityTemp. (psig) (kPa) (hours) p o p o p o p o 180° C. 1000 6894.7 6 100 10090.9 87.6 10.1 12.4 0 0 150° C. 1000 6894.7 1 100 100 100 100 0 0 0 0150° C. 500 3447.4 0.5 100 100 100 100 0 0 0 0 125° C. 100 689.5 0.5 100100 100 100 0 0 0 0

Example 10

In this example, combinatorial batch processes were evaluated using 3 wt% Pd 6 wt % Re on various carbon support materials. The reactions wererun using reaction conditions comprising 125° C., 689.5 kPa (100 psig)H₂, a reaction time of 0.5 hours, and a catalyst concentration of 2 wt %of the total feed. Results for the different carbon support materialsunder these conditions are presented in Table 5.

TABLE 5 Xylene Methyl Benzyl Conversion Selectivity Alcohol Selectivityp o p o p o Hyperion 100.0% 100.0% 100.0% 100.0% — — 07C Hyperion 100.0%100.0% 91.3% 98.9% 8.7% 1.1% 02C NoritROX 73.2% 71.7% 12.3% 10.4% 87.7%89.6% HF Norit 82.2% 81.2% 10.0% 8.5% 90.0% 91.5% Darco-LS Ceca 73.2%71.7% 12.3% 10.4% 87.7% 89.6% Pica 89.1% 87.1% 12.0% 11.7% 88.0% 88.3%Jacobi 77.6% 75.2% 2.6% 2.1% 97.4% 97.9% Nuchar 90.8% 88.0% 7.3% 7.7%92.7% 92.3%

Example 11

In this example, combinatorial batch processes were evaluated usingdifferent Group VIII metals (Pd, Pt, and Ru) on different metal oxidesupport materials (Al₂O₃, ZrO₂, and TiO₂). The reactions were run usingreaction conditions of 150° C., 4136.8 kPa (600 psig) H₂, a reactiontime of 0.5 hours, and a catalyst concentration of 2 wt % of the totalfeed. Results for the different Group VIII metals and support materialsunder these conditions are presented in Table 6.

TABLE 6 Xylene Methyl Benzyl Conversion Selectivity Alcohol SelectivityMetal Support p o p o p o 1.5% Pd Al₂O₃ ZrO₂ 100.0% 100.0% 57.1% 62.8%42.9%  37.2%  TiO₂ 100.0% 100.0% 74.4% 83.2% 25.6%  16.8%  1.5% Pt Al₂O₃32.0% 37.0% 28.8% 34.4% 71.2%  65.6%  ZrO₂ 11.3% 14.1% — — 100% 100%TiO₂ 6.1% 7.4% — — 100% 100% 1.5% Ru Al₂O₃ 40.5% 25.4% — — 100% 100%ZrO₂ 39.5% 37.8% — — 100% 100% TiO₂ 33.5% 32.5% — — 100% 100% Al₂O₃71.8% 89.6% — — 100% 100%

Example 12

In this example, combinatorial batch processes were evaluated using 3 wt% Pd and 6 wt % Re on various types of support materials. The reactionswere run using reaction conditions of 150° C., 3447.4 kPa (500 psig) H₂,a reaction time of 0.5 hours, and a catalyst concentration of 2 wt % ofthe total feed. Results for the different carbon support materials underthese conditions are presented in Table 7.

TABLE 7 Xylene Methyl Benzyl Conversion Selectivity Alcohol Selectivityp o p o p o Nb₂O₅ 93.4% 96.3% 80.4% 86.5% 19.6% 13.5% ZrO₂ 100.0% 100.0%15.3% 19.3% 84.7% 80.7% ZSM-5 73.5% 75.9% 4.9% 2.3% 95.1% 97.7% Silica82.8% 85.2% 0.8% 2.4% 99.2% 97.6% Al₂O₃ 100.0% 100.0% 12.9% 17.7% 87.1%82.3% TiO₂ 100.0% 100.0% 16.6% 17.9% 83.4% 82.1%

Example 13

In this example, combinatorial batch processes were evaluated using 3 wt% Pd, 0.75 wt % Pd, and 0.25 wt % Pd with and without Re. The reactionswere run using reaction conditions comprising 125° C., 1723.7 kPa (250psig) H₂, and a reaction time of 30 minutes at 2.8 wt % catalyst loadingon total feed. Results for this example are presented in FIG. 12 .

Example 14

In this example, combinatorial batch processes were evaluated using 0.1wt % Pd and varying amounts of Re (0.25 wt %, 1.5 wt %, and 6 wt %). Thereactions were run using reaction conditions comprising 125° C., 1723.7kPa (250 psig) H₂, and a reaction time of 30 minutes at 2.8 wt %catalyst loading on total feed. Results for this example are presentedin FIG. 13 .

Example 15

In this example, combinatorial batch processes were evaluated using 0.25wt % Pd and varying amounts of Re (0.25 wt %, 1.5 wt %, and 6 wt %). Thereactions were run using reaction conditions comprising 125° C., 1723.7kPa (250 psig) H₂, and a reaction time of 30 minutes at 2.8 wt %catalyst loading on total feed. Results for this example are presentedin FIG. 14 .

Example 16

In this example, a flow reactor was used to evaluate performance of ahydrogenation catalyst comprising Pd, Re, and an alumina supportmaterial (BASF-AL3945). The catalyst comprised 3 wt % Pd and 6 wt % Re.The reaction was run using reaction conditions comprising 180° C.,3102.6 kPa (450 psig) H₂, at 1.59 hr⁻¹ WHSV for a time on stream of over120 hours. Results for this example are presented in FIG. 15 .

Example 17

In this example, a flow reactor was used to evaluate performance of ahydrogenation catalyst comprising 0.25 wt % Pd on carbon supportmaterial (namely, Hyp07C). The reaction was run using reactionconditions comprising 180° C., 3102.6 kPa (450 psig) H₂, at 1.32 hr⁻¹WHSV for a time on stream of over 200 hours. Results for this exampleare presented in FIG. 16 .

Example 18

In this example, a flow reactor was used to evaluate performance of ahydrogenation catalyst comprising Pd, Re, and a carbon support material(namely, Hyp07C). The catalyst comprised 0.25 w % Pd and 0.5 wt % Re.The reaction was run using reaction conditions comprising 180° C.,3102.6 kPa (450 psig) H₂, at 1.32 hr⁻¹ WHSV for a time on stream of over250 hours. Results for this example are presented in FIG. 17 .

Example 19

In this example, a flow reactor was used to evaluate performance of ahydrogenation catalyst comprising Pd, Re, and a carbon support material(namely, Hyp07C). The catalyst comprised 0.1 w % Pd and 0.2 wt % Re. Thereaction was run using reaction conditions comprising 180° C., 6894.8kPa (1000 psig) H₂, at 2.32 hr⁻¹ WHSV for a time on stream of over 200hours. Results for this example are presented in FIG. 18 .

Example 20

In this example, a flow reactor was used to evaluate performance of ahydrogenation catalyst comprising 0.25 wt % Pd on carbon supportmaterial (namely, Hyp07C. The reaction was run using reaction conditionscomprising 180° C., 3102.6 kPa (450 psig) H₂, at 1.32 hr⁻¹ WHSV for atime on stream of over 600 hours. Results for this example are presentedin FIG. 19 .

Example 21

In this example, a flow reactor was used to evaluate performance of ahydrogenation catalyst comprising Pd, Re, and an alumina supportmaterial (BASF-AL3945). The catalyst comprised 0.5 wt % Pd and 1 wt %Re. The reaction was run using reaction conditions comprising 180° C.,3102.6 kPa (450 psig) H₂, at 1.32 hr⁻¹ WHSV for a time on stream of over400 hours. Results for this example are presented in FIG. 20 .

In view of the many possible embodiments to which the principles of thepresent disclosure may be applied, it should be recognized that theillustrated embodiments are only preferred examples and should not betaken as limiting the scope of the disclosure. Rather, the scope isdefined by the following claims. We therefore claim as our invention allthat comes within the scope and spirit of these claims.

We claim:
 1. An apparatus for producing p-xylene, comprising: a gasfermentation bioreactor housing a C1-fixing bacteria to produce ethanolin fluid communication with an oxidation reactor housing an oxidationcatalyst to convert the ethanol to acetaldehyde; the oxidation reactorin fluid communication with a dimerization reactor comprising adimerization catalyst to convert the acetaldehyde to 2-butenal; thedimerization reactor in fluid communication with a cyclization reactorcomprising a cyclization catalyst to convert the 2-butenal too-methylbenzaldehyde and/or p-methylbenzaldehyde; the cyclizationreactor in fluid communication with a hydrogenation reactor comprising ahydrogenation catalyst to convert the o-methylbenzaldehyde and/orp-methylbenzaldehyde to saturated cyclic compounds and non-equilibriummixture of xylenes; the hydrogenation reactor in fluid communicationwith a first fractionation zone that separates the saturated cycliccompounds from the non-equilibrium mixture of xylenes; the firstfractionation zone in fluid communication with a second fractionationzone that separates o-xylene from the non-equilibrium mixture ofxylenes; and the second fractionation zone in fluid communication with afirst crystallizer that separates p-xylene from the non-equilibriummixture of xylenes.
 2. The apparatus of claim 1, further comprising asecond crystallizer in fluid communication with the first crystallizer,wherein the second crystallizer comprises a p-xylene product outlet andan o-xylene outlet.
 3. The apparatus of claim 2, further comprising acatalytic liquid phase oxidation reactor in fluid communication with thesecond crystallizer.
 4. The apparatus of claim 3, wherein the fluidcommunication between the second crystallizer and the catalytic liquidphase oxidation reactor comprises a polymer grade p-xylene productconduit.
 5. The apparatus of claim 3, wherein the catalytic liquid phaseoxidation reactor comprises a terephthalic acid outlet.
 6. The apparatusof claim 1, further comprising a catalytic liquid phase oxidationreactor in fluid communication with the first crystallizer.
 7. Theapparatus of claim 6, wherein the fluid communication between the firstcrystallizer and the catalytic liquid phase oxidation reactor comprisesa polymer grade p-xylene product conduit.
 8. The apparatus of claim 6,wherein the catalytic liquid phase oxidation reactor comprises aterephthalic acid outlet.
 9. The apparatus of claim 1, wherein the C1fixing bacteria is in a liquid nutrient medium.
 10. The apparatus ofclaim 1, wherein the gas fermentation bioreactor comprises an industrialgas inlet, or a syngas inlet, or a hydrogen gas inlet, or anycombination thereof.
 11. The apparatus of claim 1, wherein the secondfractionation zone comprises a chemical grade o-xylene outlet.
 12. Theapparatus of claim 1, wherein the first crystallizer comprises a xyleneproduct outlet.
 13. The apparatus of claim 1, wherein the firstcrystallizer comprises a p-xylene product outlet and an o-xylene outlet.14. The apparatus of claim 1, wherein the first fractionation zonecomprises a benzene-enriched overhead outlet, adimethylcyclohexane-enriched bottoms outlet, and a remainder outlet. 15.The apparatus of claim 1, wherein the first crystallizer is sizedaccording to the size of the gas fermentation bioreactor.
 16. Theapparatus of claim 1, wherein: the fluid communication between the gasfermentation bioreactor and the oxidation reactor comprises an ethanolconduit; the fluid communication between the oxidation reactor and thedimerization reactor comprises an acetaldehyde conduit; the fluidcommunication between the dimerization reactor and the cyclizationreactor comprises an o-methylbenzaldehyde and/or p-methylbenzaldehydeconduit; the fluid communication between the cyclization reactor and thehydrogenation reactor comprises a 2-butenal conduit; the fluidcommunication between the hydrogenation reactor and the first fractionzone comprises a xylene product conduit; the fluid communication betweenthe first fraction zone and the second fractionation zone comprises aremainder conduit; and the fluid communication between the secondfraction zone and the crystallizer comprises an overhead xylene productconduit.
 17. The apparatus of claim 1, wherein the hydrogenationcatalyst comprises a first Group VIII metal deposited on a supportmaterial.
 18. The apparatus of claim 17, wherein the hydrogenationcatalyst further comprises a second Group VIII metal, a modifiercomponent, or a combination thereof and wherein the second Group VIIImetal is not the same as the first Group VIII metal.
 19. The apparatusof claim 1, wherein two or more of the dimerization reactor, thecyclization reactor, and the hydrogenation reactor, are housed within asingle vessel.
 20. The apparatus of claim 1, wherein the gasfermentation bioreactor comprises the C1-fixing bacteria in a liquidmedium and the hydrogenation catalyst comprises a first Group VIII metaldeposited on a support material.
 21. An apparatus for producingp-xylene, comprising: a gas fermentation bioreactor in fluidcommunication with an oxidation reactor; the oxidation reactor in fluidcommunication with a dimerization reactor; the dimerization reactor influid communication with a cyclization reactor; the cyclization reactorin fluid communication with a hydrogenation reactor; the hydrogenationreactor in fluid communication with a first fractionation zone; thefirst fractionation zone in fluid communication with a secondfractionation zone; and the second fractionation zone in fluidcommunication with a first crystallizer; wherein the gas fermentationbioreactor supplies ethanol to the oxidation reactor, and wherein thefirst crystallizer is used to isolate p-xylene from a non-equilibriummixture of xylenes.